Combined olefin and oxygenate conversion for aromatics production

ABSTRACT

Systems and methods are provided for inclusion of olefins in the reaction environment for an oxygenate conversion process. For conversion processes involving a metal-promoted zeolitic catalyst, addition of olefins to an oxygenate feed can reduce or minimize loss of aromatic selectivity as the catalyst is exposed to the feed. Additionally or alternately, for conversion processes involving a zeolitic catalyst including a zeolite other than an MFI framework type zeolite, addition of olefins to an oxygenate feed can reduce or minimize loss of activity for oxygenate conversion as the catalyst is exposed to the feed.

CROSS-REFERENCE TO RELATED APPLICATIONS

This application claims the benefit of U.S. Provisional Application No. 62/431,020, filed on Dec. 7, 2016, the entire contents of which are incorporated herein by reference.

FIELD

This invention relates to integrated processes for forming naphtha boiling range products, including aromatics, by conversion of oxygenates and olefins.

BACKGROUND

A variety of industrial processes are known for conversion of low boiling carbon-containing compounds to higher value products. For example, methanol to gasoline (MTG) is a commercial process that produces gasoline from methanol using ZSM-5 catalysts. In the MTG process, methanol is first dehydrated to dimethyl ether. The methanol and/or dimethyl ether then react in a series of reactions that result in formation of aromatic, paraffinic, and olefinic compounds. The resulting product consists of liquefied petroleum gas (LPG) and a high-quality gasoline comprised of aromatics, paraffins, and olefins. The typical MTG hydrocarbon product consists of about 40-50% aromatics plus olefins and about 50-60% paraffins.

U.S. Pat. No. 3,894,104 describes a method for converting oxygenates to aromatics using zeolite catalysts impregnated with a transition metal. Yields of aromatics relative to the total hydrocarbon product are reported to be as high as about 58% with a corresponding total C₅+ yield as high as about 73%.

U.S. Patent Application Publication 2013/0281753 describes a phosphorous modified zeolite catalyst. The phosphorous modification reduces the change in alpha value for the catalyst after the catalyst is exposed to an environment containing steam. The phosphorous modified catalysts are described as being suitable, for example, for conversion of methanol to gasoline boiling range compounds.

U.S. Patent Application Publications 2015/0174561, 2015/0174562, and 2015/0174563 describe catalysts for conversion of oxygenates to aromatics. The catalysts include a zeolite, such as an WI or MEL framework structure zeolite, with a supported Group 12 metal on the catalyst.

U.S. Pat. No. 9,090,525 describes conversion of oxygenates in the presence of a zeolitic catalyst to form naphtha boiling range compounds with increased octane. A portion of the naphtha boiling range olefins from an initial conversion product are recycled to the oxygenate conversion process to allow for formation of heavier naphtha boiling range compounds, including aromatics.

SUMMARY

In some aspects, a method for forming a naphtha composition is provided. The method can include exposing a feed comprising oxygenates (such as methanol) and olefins to a conversion catalyst. Examples of effective conversion conditions can include an average reaction temperature of about 300° C. to about 550° C., a total pressure of about 10 psig (˜70 kPag) to about 400 psig (˜2700 kPag), and/or a WHSV of 0.1 hr⁻¹ to 20.0 hr⁻¹. Exposing the feed to the conversion catalyst can result in formation of a converted effluent comprising a naphtha boiling range fraction. The naphtha boiling range fraction can have an octane rating of at least 80 and/or can comprise less than 6.0 wt % combined of CO, CO₂, and CH₄ relative to a total weight of hydrocarbons in the converted effluent. The feed can have a molar ratio of oxygenates to olefins of about 1 to about 20. The conversion catalyst can comprise at least 10 wt % of a zeolite having MFI framework structure, the zeolite optionally having a silicon to aluminum ratio of 10 to 200 and an Alpha value of at least 5. The conversion catalyst can further comprise 0.1 wt % to 3.0 wt % of a transition metal supported on the conversion catalyst, such as zinc supported on the conversion catalyst.

Optionally, the naphtha boiling range fraction can have an octane rating of at least 90 (or at least 93, or at least 97) and/or at least about 40 wt % aromatics relative to a weight of the naphtha boiling range fraction. Optionally, the average reaction temperature can be at least about 400° C., or at least about 450° C.

In some aspects, a method for forming a naphtha composition is provided. The method can include exposing a feed comprising oxygenates (such as methanol) and olefins to a conversion catalyst. Examples of effective conversion conditions can include an average reaction temperature of about 300° C. to about 550° C., a total pressure of about 10 psig (˜70 kPag) to about 400 psig (˜2700 kPag), and/or a WHSV of 0.1 hr⁻¹ to 20.0 hr⁻¹. Exposing the feed to the conversion catalyst can result in formation of a converted effluent comprising a naphtha boiling range fraction and optionally further comprising at least about 30 wt % olefins and/or less than 15 wt % oxygenate relative to a total weight of hydrocarbons in the converted effluent. The feed can have a molar ratio of oxygenates to olefins of 1 to 20. The conversion catalyst can comprise at least 10 wt % of a 10-member ring or 12-member ring zeolite having a framework structure different from MFI framework structure. The zeolite can optionally have a silicon to aluminum ratio of 10 to 200 and/or an Alpha value of at least 5. The conversion catalyst can optionally further comprise an average catalyst exposure time of 25 grams to 200 grams of oxygenate per gram of catalyst. The conversion catalyst can optionally further comprise 0.1 wt % to 3.0 wt % of a transition metal supported on the conversion catalyst, such as Zn.

Optionally, the conversion catalyst can comprise an average catalyst exposure time of 50 grams to 200 grams of methanol per gram of catalyst, (or 25 grams to 180 grams, or 50 grams to 180 grams, or 50 grams to 150 grams, or 100 grams to 200 grams). Optionally, the conversion catalyst can comprise at least 10 wt % of a zeolite having a framework structure of MRE (ZSM-48), MTW, TON, MTT, MFS, or a combination thereof.

In an optional aspect, exposing the feed comprising oxygenates to a conversion catalyst comprises exposing the feed comprising oxygenate to the conversion catalyst in a fluidized bed, a moving bed, a riser reactor, or a combination thereof, the conversion catalyst being withdrawn and regenerated at a rate corresponding to regeneration of 0.3 wt % to 3.0 wt % of catalyst per 1 g of oxygenate exposed to a g of conversion catalyst (optionally 1.5 wt % to 3.0 wt %).

In some aspects, a conversion catalyst can further comprise phosphorus supported on the conversion catalyst, a molar ratio of phosphorus to zinc on the conversion catalyst optionally being 1.5 to 3.0. Additionally or alternately, the feed comprising oxygenates and olefins comprises a first feedstock comprising at least a portion of the oxygenates and a second feedstock comprising at least a portion of the olefins, the first feedstock and the second feedstock being combined after entering a reactor containing the conversion catalyst.

In some optional aspects, a) the feed can comprise about 30 wt % to about 95 wt % of oxygenates, about 5 wt % to about 40 wt % of olefins, or a combination thereof; b) the feed can comprise at least about 20 wt % to about 60 wt % of components different from oxygenates and olefins, or about 40 wt % to 60 wt %; or c) a combination of a) and b).

In still another aspect, an oxygenate conversion effluent is provided. The oxygenate conversion effluent can include, relative to a total weight of hydrocarbons in the conversion effluent, at least 40 wt % aromatics, less than 6.0 wt % combined of CO, CO₂, and CH₄, and less than 10 wt % olefins. A naphtha boiling range portion of the conversion effluent can have an octane rating of at least 90. Optionally, less than 10 wt % of the aromatics comprise C₁₀ aromatics relative to a total weight of the aromatics, and/or less than 10 wt % of the C₁₀ aromatics comprise durene relative to a total weight of the C₁₀ aromatics. Optionally, the oxygenate conversion effluent can comprise less than 5.0 wt % combined of CO, CO₂, and CH₄. Optionally, less than 5 wt % of the C₁₀ aromatics comprise durene relative to a total weight of the C₁₀ aromatics, or a combination thereof.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 schematically shows an example of a reaction system for conversion of oxygenates to olefins.

FIG. 2 shows results from conversion of methanol in the presence of a variety of zeolitic catalysts.

FIG. 3 shows relative yields from conversion of methanol and methanol plus 1-pentene in the presence of a 1 wt % Zn-ZSM-5 catalyst.

FIG. 4 shows results from conversion of methanol at various temperatures in the presence of a ZSM-5 catalyst.

FIG. 5 shows results from conversion of methanol and 1-pentene at various temperatures in the presence of a ZSM-5 catalyst.

FIG. 6 shows relative yields from conversion of methanol and methanol plus 1-pentene in the presence of a ZSM-5 catalyst.

FIG. 7 shows another type of relative yield analysis from conversion of methanol and methanol plus 1-pentene in the presence of a ZSM-5 catalyst.

FIG. 8 shows relative aromatic yields from conversion of methanol and methanol plus 1-pentene in the presence of a ZSM-5 catalyst.

FIG. 9 shows relative aromatic yields from conversion of methanol and methanol plus 1-pentene in the presence of a 1 wt % Zn-ZSM-5 catalyst.

FIG. 10 shows results from conversion of methanol in the presence of a ZSM-48 catalyst.

FIG. 11 shows results from conversion of methanol plus 1-pentene in the presence of a ZSM-48 catalyst.

DETAILED DESCRIPTION

In various aspects, systems and methods are provided for conversion of a combined feed of oxygenates (such as methanol or dimethyl ether) and olefins to a high octane naphtha boiling range product. For conversion processes involving a metal-promoted zeolitic catalyst, addition of olefins to an oxygenate feed can reduce or minimize loss of aromatic selectivity as the catalyst is exposed to the feed. Additionally or alternately, for conversion processes involving a zeolitic catalyst other than an MFI framework type, addition of olefins to an oxygenate feed can reduce or minimize loss of activity for oxygenate conversion as the catalyst is exposed to the feed.

Natural gas, coal, and/or biomass are becoming increasingly important sources of carbon for use in production of fuel and/or lubricant products. A first step in conversion of carbon from a natural gas, coal, and/or biomass source can be a conversion of methane to methanol. Once methanol is formed, various fixed bed, fluid bed, and moving bed processes can be used to convert methanol to higher value products, such as fuels, aromatics, and/or olefins. Such processes can use zeolitic catalysts, such as MFI framework (ZSM-5) zeolitic catalysts.

Some difficulties with conversion of methanol to naphtha boiling range products (such as aromatics) for use as gasoline can be related to the tendency for the zeolitic catalyst to deactivate relatively quickly. Even relatively small exposures of feed to a zeolitic catalyst can result in loss of aromatic selectivity, with a corresponding increase in formation of lower value paraffins. For zeolitic frameworks other than MFI, the catalyst deactivation can also impact the general ability of the catalyst to convert oxygenates within a feed.

One option for increasing the aromatic selectivity of a zeolitic catalyst can be to add a supported transition metal on the catalyst to promote aromatic formation. Zinc is an example of a suitable transition metal to improve aromatic selectivity. While zinc can be effective for increasing the initial aromatic selectivity of a zeolitic catalyst, such as MFI framework zeolitic catalyst, exposure to an oxygenate feed can cause the catalyst to lose aromatic selectivity due to conversion of oxygenates to carbon oxides and paraffins.

It has been unexpectedly discovered that addition of olefins to an oxygenate feed can reduce or minimize loss of catalyst activity and/or selectivity during exposure to the oxygenate feed. The olefins can correspond to any convenient type of C₂-C₆ olefin. In some aspects, the olefins can correspond to olefins generated during the oxygenate conversion process. In such aspects, a portion of the effluent from the conversion process can be recycled to provide olefins for the feed. In other aspects, the olefins can be derived from any other convenient source. The olefin feed can optionally include compounds that act as inerts or act as a diluent in the conversion process. For example, a stream of “waste” olefins having an olefin content of 5 vol % to 20 vol % can be suitable as a source of olefins, so long as the other components of the “waste” olefins stream are compatible with the conversion process. For example, the other components of the olefin stream can correspond to inert gases such as N₂, carbon oxides, paraffins, and/or other gases that have low reactivity under the conversion conditions. Water can also be present, although it can be preferable for water to correspond to 20 vol % or less of the total feed, or 10 vol % or less.

In this discussion, octane rating is defined as (RON+MON)/2, where RON is research octane number and MON is motor octane number. For values reported in the examples below, RON and MON values were determined based on a published model that determines octane ratings for a blend of components based to determine a blended octane. The model is described at Ind Eng Chem Res 2006, 45, 337-345. The model is believed to correlate with experimentally determined values. In the claims below, Research Octane Number (RON) is determined according to ASTM D2699. Motor Octane Number (MON) is determined according to ASTM D2700.

In this discussion, the naphtha boiling range is defined as 50° F. (˜10° C., roughly corresponding to the lowest boiling point of a pentane isomer) to 350° F. (177° C.). The distillate fuel boiling range, is defined as 350° F. (177° C.) to 700° F. (371° C.). Compounds (C⁴⁻) with a boiling point below the naphtha boiling range can be referred to as light ends. It is noted that due to practical consideration during fractionation (or other boiling point based separation) of hydrocarbon-like fractions, a fuel fraction formed according to the methods described herein may have T5 and T95 distillation points corresponding to the above values (or T10 and T90 distillation points), as opposed to having initial/final boiling points corresponding to the above values. While various methods are available for determining boiling point information for a given sample, for the claims below ASTM D86 is a suitable method for determining distillation points (including fractional weight distillation points) for a composition.

Feedstocks and Products—Oxygenate Conversion

In various aspects, catalysts described herein can be used for conversion of oxygenate feeds to aromatics and/or olefins products, such as oxygenates containing at least one C₁-C₄ alkyl group and/or other oxygenates. Examples of suitable oxygenates include feeds containing methanol, dimethyl ether, C₁-C₄ alcohols, ethers with C₁-C₄ alkyl chains, including both asymmetric ethers containing C₁-C₄ alkyl chains (such as methyl ethyl ether, propyl butyl ether, or methyl propyl ether) and symmetric ethers (such as diethyl ether, dipropyl ether, or dibutyl ether), or combinations thereof. It is noted that oxygenates containing at least one C₁-C₄ alkyl group are intended to explicitly identify oxygenates having alkyl groups containing about 4 carbons or less. Preferably the oxygenate feed can include at least about 30 wt % of one or more suitable oxygenates, or at least about 50 wt %, or at least about 75 wt %, or at least about 90 wt %, or at least about 95 wt %. Additionally or alternately, the oxygenate feed can include at least about 50 wt % methanol, such as at least about 75 wt % methanol, or at least about 90 wt % methanol, or at least about 95 wt % methanol. In particular, the oxygenate feed can include 30 wt % to 100 wt % of oxygenate (or methanol), or 50 wt % to 95 wt %, or 75 wt % to 100 wt %, or 75 wt % to 95 wt %. The oxygenate feed can be derived from any convenient source. For example, the oxygenate feed can be formed by reforming of hydrocarbons in a natural gas feed to form synthesis gas (H₂, CO, CO₂), and then using the synthesis gas to form methanol (or other alcohols). As another example, a suitable oxygenate feed can include methanol, dimethyl ether, or a combination thereof as the oxygenate.

In addition to oxygenates, the feed can also include olefins. In this discussion, the olefins included as part of the feed can correspond to aliphatic olefins that contain 6 carbons or less, so that the olefins are suitable for formation of naphtha boiling range compounds. The olefin portion of the feed can be mixed with the oxygenates prior to entering a reactor for performing oxygenate conversion, or a plurality of streams containing oxygenates and/or olefins can be mixed within a conversion reactor. The feed can include about 5 wt % to about 40 wt % of olefins (i.e., olefins containing 6 carbons or less), or about 5 wt % to about 30 wt %, or about 10 wt % to about 40 wt %, or about 10 wt % to about 30 wt %. Without being bound by any particular theory, it is believed that olefins can compete effectively for active sites on a zeolitic catalyst that have high activity for conversion of oxygenates to paraffins or carbon oxides, such as conversion of methanol to methane. In order to enable this competitive effect so that olefins suppress undesirable activity, the molar ratio of oxygenates to olefins can be 20 or less, or 10 or less, or 6.0 or less, or 4.0 or less, such as down to a molar ratio of about 1.0. It is noted that the weight percent of olefins in the feed can be dependent on the nature of the olefins. For example, if a C₅ olefin is used as the olefin with a methanol-containing feed, the wt % of olefin required to achieve a desired molar ratio of olefin to oxygenate will be relatively high due to the much larger molecular weight of a C₅ alkene.

In addition to oxygenates and olefins, a feed can also include diluents, such as water (in the form of steam), nitrogen or other inert gases, and/or paraffins or other non-reactive hydrocarbons. In some aspects, the source of olefins can correspond to a low purity source of olefins, so that the source of olefins corresponds to 20 wt % or less of olefins. In some aspects, the portion of the feed corresponding to components different from oxygenates and olefins can correspond to 1 wt % to 60 wt % of the feed, or 1 wt % to 25 wt %, or about 10 wt % to about 30 wt %, or about 20 wt % to about 60 wt %. Optionally, the feed can substantially correspond to oxygenates and olefins, so that the content of components different from oxygenates and olefins is 1 wt % or less (such as down to 0 wt %).

In some aspects, such as aspects related to oxygenate conversion using an MFI or MEL framework catalyst, the yield of aromatics relative to the total hydrocarbon product can be about 35 wt % to about 60 wt %, or about 38 wt % to about 60 wt %, or about 40 wt % to about 52 wt %, or about 38 wt % to about 45 wt %. For example, the yield of aromatics relative to the total hydrocarbon product can be at least about 35 wt %, or at least about 38 wt %, or at least about 40 wt %, or at least about 45 wt %. Additionally or alternately, the yield of aromatics relative to the total hydrocarbon product can be about 60 wt % or less, or about 55 wt % or less, or about 52 wt % or less, or about 50 wt % or less. In various aspects, the yield of olefins relative to the total hydrocarbon product can be about 2.0 wt % to about 30 wt %, or about 2.0 wt % to 25 wt %, or about 5.0 wt % to about 20 wt %, or about 10 wt % to about 20 wt %. For example, the yield of olefins relative to the total hydrocarbon product can be at least about 2.0 wt %, or at least about 5.0 wt %, or at least about 10 wt %. Additionally or alternately, the yield of olefins relative to the total hydrocarbon product can be about 30 wt % or less, or about 25 wt % or less, or about 20 wt % or less. In various aspects, the yield of paraffins relative to the total hydrocarbon product can be about 20 wt % to about 45 wt %, or about 20 wt % to about 35 wt %, or about 25 wt % to about 45 wt %, or about 25 wt % to about 40 wt %. For example, the yield of paraffins relative to the total hydrocarbon product can be at least about 20 wt %, or at least about 25 wt %, or at least about 30 wt % and/or the yield of paraffins relative to the total hydrocarbon product can be about 45 wt % or less, or about 40 wt % or less, or about 35 wt % or less. In the claims below, the relative amounts of paraffins, olefins, and aromatics in a sample can be determined based on ASTM D6839. For the paraffins and olefins generated during oxygenate conversion, at least 50 wt % of the olefins can correspond to C₃ and C₄ olefins and/or at least 50 wt % of the paraffins can correspond to C₃ and C₄ paraffins. Additionally or alternately, less than 10 wt % of the paraffins can correspond to C₁ paraffins (methane).

In some aspects, such as aspects related to oxygenate conversion using an MRE framework catalyst, the yield of aromatics relative to the total hydrocarbon product can be about 5 wt % to about 30 wt %, or about 10 wt % to about 30 wt %, or about 10 wt % to about 25 wt %, or about 5 wt % to about 20 wt %. For example, the yield of aromatics relative to the total hydrocarbon product can be at least about 5 wt %, or at least about 10 wt %, or at least about 15 wt %. Additionally or alternately, the yield of aromatics relative to the total hydrocarbon product can be about 30 wt % or less, or about 25 wt % or less, or about 20 wt % or less. In various aspects, the yield of olefins relative to the total hydrocarbon product can be about 20 wt % to about 60 wt %, or about 25 wt % to 60 wt %, or about 20 wt % to about 40 wt %, or about 25 wt % to about 50 wt %. For example, the yield of olefins relative to the total hydrocarbon product can be at least about 20 wt %, or at least about 25 wt %, or at least about 30 wt %. Additionally or alternately, the yield of olefins relative to the total hydrocarbon product can be about 60 wt % or less, or about 50 wt % or less, or about 40 wt % or less. In various aspects, the yield of paraffins relative to the total hydrocarbon product can be about 20 wt % to about 50 wt %, or about 20 wt % to about 35 wt %, or about 25 wt % to about 45 wt %, or about 25 wt % to about 40 wt %. For example, the yield of paraffins relative to the total hydrocarbon product can be at least about 20 wt %, or at least about 25 wt %, or at least about 30 wt % and/or the yield of paraffins relative to the total hydrocarbon product can be about 50 wt % or less, or about 45 wt % or less, or about 40 wt % or less, or about 35 wt % or less. For the paraffins and olefins generated during oxygenate conversion, at least 50 wt % of the olefins can correspond to C₃ and C₄ olefins and/or at least 50 wt % of the paraffins can correspond to C₃ and C₄ paraffins. Additionally or alternately, less than 10 wt % of the paraffins can correspond to C₁ paraffins (methane).

The total hydrocarbon product in the conversion effluent can include a naphtha boiling range portion, a distillate fuel boiling range portion, and a light ends portion. Optionally but preferably, the conversion effluent can include less than 1.0 wt % of compounds boiling above the distillate fuel boiling range (371° C+), such as having a final boiling point of 371° C. or less. In various aspects, the selectivity for forming/yield of a naphtha boiling range portion can be at least about 35 wt % and/or about 75 wt % or less. For example, the selectivity for forming/yield of a naphtha boiling range portion can be about 35 wt % to 75 wt %, or 40 wt % to 65 wt %, or 40 wt % to 60 wt %, or 45 wt % to 70 wt %.

The naphtha boiling range portion formed from a conversion process can have an octane rating of at least 80, or at least 90, or at least 95, or at least 97, or at least 100, or at least 102, or at least 105, such as up to 110. In particular, in aspects involving an MFI or MEL framework catalyst, the octane rating can be 80 to 110, or 95 to 110, or 97 to 110, or 100 to 110. Additionally or alternately, in aspects involving a MRE framework catalyst, the octane rating can be 80 to 97 or 90 to 97. As defined above, the octane rating is corresponds to (RON+MON)/2).

The conversion conditions can also result in generation of CO and/or CO₂. In some aspects, the amount of combined CO, CO₂, and CH₄ can correspond to about 6.0 wt % or less of the total hydrocarbon product in a conversion effluent, or about 5.0 wt % or less. In this discussion and the claims below, the amounts of CO and CO₂ in a conversion effluent are included when determining the amount of the total hydrocarbon product (such as the weight of the total hydrocarbon product).

Suitable and/or effective conditions for performing a conversion reaction can include average reaction temperatures of about 300° C. to about 550° C. (or about 350° C. to about 550° C., or about 400° C. to about 500° C.), total pressures between about 10 psig (˜70 kPag) to about 400 psig (˜2700 kPag), or about 50 psig (˜350 kPag) to about 350 psig (˜2400 kPag), or about 100 psig (˜700 kPag) to about 300 psig (˜2100 kPag), and an oxygenate space velocity between about 0.1 h⁻¹¹ to about 10 h⁻¹ based on weight of oxygenate relative to weight of catalyst. For example, the average reaction temperature can be at least about 300° C., or at least about 350° C., or at least about 400° C., or at least about 450° C. Additionally or alternately, the average reaction temperature can be about 550° C. or less, or about 500° C. or less, or about 450° C. or less, or about 400° C. or less. In this discussion, average reaction temperature is defined as the average of the temperature at the reactor inlet and the temperature at the reactor outlet for the reactor where the conversion reaction is performed. As another example, the total pressure can be at least about 70 kPag, or at least about 350 kPag, or at least about 500 kPag, or at least about 700 kPag, or at least about 1000 kPag. Additionally or alternately, the total pressure can be about 3000 kPag or less, or about 2700 kPag or less, or about 2400 kPag or less, or about 2100 kPag or less.

Optionally, a portion of the conversion effluent can be recycled for inclusion as part of the feed to the conversion reactor. For example, at least a portion of the light ends from the conversion effluent can be recycled as part of the feed. The recycled portion of the light ends can correspond to any convenient amount, such as 25 wt % to 75 wt % of the light ends. Recycling of light ends can provide olefins, which can serve as an additional reactant in the conversion reaction, as well as providing a mechanism for temperature control.

Various types of reactors can provide a suitable configuration for performing a conversion reaction. Suitable reactors can include fixed bed reactors (such as trickle bed reactors), moving bed reactors (such as riser reactors), and fluidized bed reactors. It is noted that the activity and/or selectivity of a catalyst for oxygenate conversion can vary as the catalyst is exposed to increasing amounts of oxygenate feed. This modification of the catalyst activity is believed to be due to the formation of coke on the catalyst. When oxygenate conversion is performed in a fixed bed reactor, calculating the average catalyst exposure time can be straightforward, as the amount of oxygenate introduced into the reactor can be compared with the amount of conversion catalyst in the reactor. This can be used to calculate an average catalyst exposure time as a ratio of the grams of oxygenate (such as methanol) exposed to the catalyst divided by the grams of catalyst.

The modification of the catalyst activity and/or selectivity with increasing average catalyst exposure time can be reversed at least in part by regenerating the catalyst. In some aspects, a full regeneration can be performed on a catalyst, so that the average amount of coke present on the regenerated catalyst is less than 0.1 wt %. In other aspects, a partial regeneration can be performed, so that the average amount of coke present on the regenerated catalyst after regeneration is greater than 0.1 wt %. The average amount of coke present on a catalyst sample can be readily determined by thermogravimetric analysis.

In aspects where a catalyst can be withdrawn from the reactor for regeneration and recycle during operation of the reactor, such as a moving bed reactor and/or fluidized bed reactor, catalyst can be withdrawn and replaced with make-up (fresh) and/or regenerated catalyst. It is noted that withdrawing catalyst from the reactor for regeneration is distinct from removing catalyst entirely from the reaction system and replacing the removed catalyst with fresh make-up catalyst. In this discussion, when full regeneration is performed on a catalyst (less than 0.1 wt % average coke remaining on the regenerated catalyst), the average catalyst exposure time for the regenerated catalyst is defined to be zero for purposes of determining average catalyst exposure time for catalyst within the reactor. In such aspects when full regeneration is being performed, the average catalyst exposure time for catalyst being exposed to oxygenate can be determined based on a) the flow rate of oxygenate into the reactor relative to the amount of catalyst in the reactor, and b) the average residence time of the catalyst in the reactor. These values can allow for a determination of the average grams of oxygenate per gram of catalyst in the reactor (i.e., the average catalyst exposure time).

In a moving bed reactor, the residence time for catalyst can correspond to the amount of time required for a catalyst particle to travel the length of the bed to the exit, based on the average velocity of the moving bed. As an example, the flow of methanol into a moving bed reactor can correspond to a space velocity of 1.0 hr⁻¹, which means 1 g of methanol per g of catalyst per hour. In such an example, if the average residence time for catalyst in the reactor is 48 hours (based on the average velocity of the moving bed relative to the size of the bed), then the average catalyst exposure time for catalyst in the moving bed would be 24 g of methanol per g of catalyst. Similarly, in aspects involving a fluidized bed, the catalyst residence time can be determined based on the rate of removal of catalyst from the reactor for regeneration. The catalyst residence time can correspond to the amount of time required to remove an amount of catalyst that is equivalent to the weight of the catalyst bed. Based on that residence time, the average catalyst exposure time can be calculated in a similar manner to the calculation for a moving bed.

During a partial regeneration, a catalyst can be exposed to an oxidizing environment for removal of coke from the catalyst, but the net amount of coke remaining on the catalyst after partial regeneration can be greater than 0.1 wt %. When a partial regeneration is performed, the effective average catalyst exposure time for the catalyst after regeneration will be a value other than zero, due to the amount of remaining coke on the catalyst. When a partial regeneration is performed, the amount of coke removal can roughly scale in a linear manner with the effective average catalyst exposure time of the partially regenerated catalyst. In this discussion and the claims below, when a catalyst is partially regenerated, the average catalyst exposure time for the partially regenerated catalyst is determined by multiplying the average catalyst exposure time prior to regeneration by the wt % of coke remaining on the catalyst after partial regeneration. As an example, a hypothetical catalyst may have an exposure time of 100 g methanol per g catalyst prior to regeneration. In this example, partial regeneration is used to remove 60 wt % of the coke on the catalyst. This means that 40 wt % (or 0.4 expressed as a fraction) of the coke remains on the catalyst after regeneration. In such an example, the average catalyst exposure time for the regenerated catalyst would be 0.4×100=40 g methanol per g catalyst.

In aspects where partial regeneration is performed, the calculation for the average catalyst exposure time for catalyst in the reactor can be modified based to account for the fact that any recycled catalyst will have a non-zero initial value of catalyst exposure time. The same calculation described above can be used to determine an initial value. The non-zero catalyst exposure time for the regenerated catalyst can then be added to the initial value to determine the average catalyst exposure time within the reactor. In the example noted above, if the average catalyst exposure time for partially regenerated catalyst is 10 g methanol per g catalyst, and if the amount of average exposure within the reactor is 24 g methanol per g catalyst as calculated above, then the average catalyst exposure time for the system when using partial regeneration would be 34 g methanol per g catalyst. It is also noted that a portion of the catalyst introduced into a reactor may correspond to fresh make-up catalyst instead of partially regenerated catalyst. In such aspects, the catalyst exposure time for the catalyst introduced into the reactor can be a weighted average of the fresh make-up catalyst (zero exposure time) and the catalyst exposure time for the partially regenerated catalyst.

For a catalyst including an MFI framework zeolite, the catalyst recycle rate can be dependent on the desired products, with catalyst recycle rates that produce an average catalyst exposure time/average cycle length for catalyst in the reactor of about 1 g CH₃OH/g catalyst to about 2000 g CH₃OH/g catalyst potentially being suitable, or about 50 g CH₃OH/g catalyst to about 1000 g CH₃OH/g catalyst, or about 100 g CH₃OH/g catalyst to about 1500 g CH₃OH/g catalyst, or about 100 g CH₃OH/g catalyst to about 1000 g CH₃OH/g catalyst. The target average catalyst exposure time can be dependent on the specific nature of the catalyst and/or the desired product mix. In some aspects where shorter average catalyst exposure times are desired, the average catalyst exposure time can be about 1 g CH₃OH/g catalyst to about 200 g CH₃OH/g catalyst, or about 5 g CH₃OH/g catalyst to about 150 g CH₃OH/g catalyst, or about 1 g CH₃OH/g catalyst to about 100 g CH₃OH/g catalyst. In other aspects where longer times are desired, the average catalyst exposure time can be about 200 g CH₃OH/g catalyst to about 2000 g CH₃OH/g catalyst, or about 400 g CH₃OH/g catalyst to about 1500 g CH₃OH/g catalyst, or about 500 g CH₃OH/g catalyst to about 1000 g CH₃OH/g catalyst. The above average catalyst exposure times can be achieved, for example, by withdrawing about 0.01 wt % to about 3.0 wt % of catalyst per 1 g of methanol exposed to a g of conversion catalyst, or about 0.01 wt % to about 1.5 wt %, or about 0.1 wt % to about 3.0 wt %, or about 1.0 wt % to about 3.0 wt %. It is noted that these withdrawal rates could be modified, for example, if only a partial regeneration is performed on withdrawn catalyst. For catalysts other than MFI framework catalysts, a catalyst recycle rate can be selected to produce an average catalyst exposure time/average cycle length for catalyst in the reactor of about 25 g CH₃OH/g catalyst to about 200 g CH₃OH/g catalyst, or about 25 g CH₃OH/g catalyst to about 180 g CH₃OH/g catalyst, or about 50 g CH₃OH/g catalyst to about 180 g CH₃OH/g catalyst, or about 50 g CH₃OH/g catalyst to about 150 g CH₃OH/g catalyst, or about 25 g CH₃OH/g catalyst to about 100 g CH₃OH/g catalyst, or about 50 g CH₃OH/g catalyst to about 100 g CH₃OH/g catalyst, or about 100 g CH₃OH/g catalyst to about 180 g CH₃OH/g catalyst, or about 100 g CH₃OH/g catalyst to about 150 g CH₃OH/g catalyst. The appropriate cycle length for a catalyst including a non-MFI framework zeolite can depend on the type of zeolite.

It is noted that the oxygenate feed and/or conversion reaction environment can include water in various proportions. Conversion of oxygenates to aromatics and olefins results in production of water as a product, so the relative amounts of oxygenate (such as methanol or dimethyl ether) and water can vary within the reaction environment. Based on the temperatures present during methanol conversion, the water in the reaction environment can result in “steaming” of a catalyst. Thus, a catalyst used for conversion of oxygenates to aromatics is preferably a catalyst that substantially retains activity when steamed. Water may also be present in a feed prior to contacting the zeolite catalyst. For example, in commercial processing of methanol to form gasoline, in order to control heat release within a reactor, an initial catalyst stage may be used to convert a portion of the methanol in a feed to dimethyl ether and water prior to contacting a zeolite catalyst for forming gasoline.

Catalysts for Oxygenate Conversion

In various aspects, a transition metal-enhanced zeolite catalyst composition can be used for conversion of oxygenate feeds to naphtha boiling range fractions and olefins. In this discussion and the claims below, a zeolite is defined to refer to a crystalline material having a porous framework structure built from tetrahedra atoms connected by bridging oxygen atoms. Examples of known zeolite frameworks are given in the “Atlas of Zeolite Frameworks” published on behalf of the Structure Commission of the International Zeolite Association”, 6^(th) revised edition, Ch. Baerlocher, L. B. McCusker, D. H. Olson, eds., Elsevier, N.Y. (2007) and the corresponding web site, http://www.iza-structure.org/databases/. Under this definition, a zeolite can refer to aluminosilicates having a zeolitic framework type as well as crystalline structures containing oxides of heteroatoms different from silicon and aluminum. Such heteroatoms can include any heteroatom generally known to be suitable for inclusion in a zeolitic framework, such as gallium, boron, germanium, phosphorus, zinc, and/or other transition metals that can substitute for silicon and/or aluminum in a zeolitic framework.

A suitable zeolite can include a 10-member or 12-member ring pore channel network, such as a 1-dimensional 10-member ring pore channel or a 3-dimensional 10-member ring pore channel. Examples of suitable zeolites having a 3-dimensional 10-member ring pore channel network include zeolites having an MFI or MEL framework, such as ZSM-5 or ZSM-11. ZSM-5 is described in detail in U.S. Pat. Nos. 3,702,886 and Re. 29,948. ZSM-11 is described in detail in U.S. Pat. No. 3,709,979. Preferably, the zeolite is ZSM-5. Examples of suitable zeolites having a 1-dimensional 10-member ring pore channel network include zeolites having a MRE (ZSM-48), MTW, TON, MTT, and/or MFS framework. In some aspects, a zeolite with a 3-dimensional pore channel can be preferred for conversion of methanol, such as a zeolite with an MFI framework.

Generally, a zeolite having desired activity for methanol conversion can have a silicon to aluminum molar ratio of about 10 to about 200, or about 15 to about 100, or about 20 to about 80, or about 20 to about 40. For example, the silicon to aluminum ratio can be at least about 10, or at least about 20, or at least about 30, or at least about 40, or at least about 50, or at least about 60. Additionally or alternately, the silicon to aluminum ratio can be about 300 or less, or about 200 or less, or about 100 or less, or about 80 or less, or about 60 or less, or about 50 or less.

Typically, reducing the silicon to aluminum ratio in a zeolite will result in a zeolite with a higher acidity, and therefore higher activity for cracking of hydrocarbon or hydrocarbonaceous feeds, such as petroleum feeds. However, with respect to conversion of oxygenates to aromatics, such increased cracking activity may not be beneficial, and instead may result in increased formation of residual carbon or coke during the conversion reaction. Such residual carbon can deposit on the zeolite catalyst, leading to deactivation of the catalyst over time. Having a silicon to aluminum ratio of at least about 40, such as at least about 50 or at least about 60, can reduce or minimize the amount of additional residual carbon that is formed due to the acidic or cracking activity of a catalyst.

It is noted that the molar ratio described herein is a ratio of silicon to aluminum. If a corresponding ratio of silica to alumina were described, the corresponding ratio of silica (SiO₂) to alumina (Al₂O₃) would be twice as large, due to the presence of two aluminum atoms in each alumina stoichiometric unit. Thus, a silicon to aluminum ratio of 10 corresponds to a silica to alumina ratio of 20.

In some aspects, a zeolite in a catalyst can be present at least partly in the hydrogen form. Depending on the conditions used to synthesize the zeolite, this may correspond to converting the zeolite from, for example, the sodium form. This can readily be achieved, for example, by ion exchange to convert the zeolite to the ammonium form followed by calcination in air or an inert atmosphere at a temperature of about 400° C. to about 700° C. to convert the ammonium form to the active hydrogen form.

Additionally or alternately, a zeolitic catalyst can include and/or be enhanced by a transition metal. Preferably the transition metal is a Group 12 metal from the IUPAC periodic table (sometimes designated as Group IIB) selected from Zn, Cd, or a combination thereof. More generally, the transition metal can be any convenient transition metal selected from Groups 6-15 of the IUPAC periodic table. The transition metal can be incorporated into the zeolite/catalyst by any convenient method, such as by impregnation, by ion exchange, by mulling prior to extrusion, and/or by any other convenient method. Optionally, the transition metal incorporated into a zeolite/catalyst can correspond to two or more metals. After impregnation or ion exchange, the transition metal-enhanced catalyst can be treated in air or an inert atmosphere at a temperature of about 400° C. to about 700° C. The amount of transition metal can be expressed as a weight percentage of metal relative to the total weight of the catalyst (including any zeolite and any binder). A catalyst can include about 0.05 wt % to about 20 wt % of one or more transition metals, or about 0.1 wt % to about 10 wt %, or about 0.1 wt % to about 5 wt %, or about 0.1 wt % to about 2.0 wt %. For example, the amount of transition metal can be at least about 0.1 wt % of transition metal, or at least about 0.25 wt % of transition metal, or at least about 0.5 wt %, or at least about 0.75 wt %, or at least about 1.0 wt %. Additionally or alternately, the amount of transition metal can be about 20 wt % or less, or about 10 wt % or less, or about 5 wt % or less, or about 2.0 wt % or less, or about 1.5 wt % or less, or about 1.2 wt % or less, or about 1.1 wt % or less, or about 1.0 wt % or less.

In some optional aspects, a zeolitic catalyst can be substantially free of phosphorous. A catalyst composition that is substantially free of phosphorous can contain about 0.01 wt % of phosphorous or less, such as less than about 0.005 wt % of phosphorous, or less than about 0.001 wt % of phosphorous. A zeolitic catalyst that is substantially free of phosphorous can be substantially free of intentionally added phosphorous or substantially free of both intentionally added phosphorous as well as phosphorous present as an impurity in a reagent for forming the catalyst composition. In some aspects, a zeolitic catalyst can contain no added phosphorous, such as containing no intentionally added phosphorous and/or containing no phosphorous impurities to within the detection limits of standard methods for characterizing a reagent and/or a resulting zeolite.

Optionally, a zeolitic catalyst for methanol conversion can include added phosphorus, such as phosphorus added by impregnation, ion exchange, mulling prior to extrusion, or another convenient method. The amount of phosphorus can be related to the amount of transition metal in the catalyst composition. In some aspects, the molar ratio of phosphorus to transition metal can be 0.5 to 5.0, or 1.5 to 3.0, or 1.0 to 2.5, or 1.5 to 2.5. At higher molar ratios of phosphorus to transition metal, the phosphorus can be beneficial for maintaining a relatively stable selectivity for aromatics formation during an oxygenate conversion process. Additionally or alternately, a catalyst can include about 0.05 wt % to about 10 wt % of phosphorus, or about 0.1 wt % to about 10 wt %, or about 0.1 wt % to about 5 wt %, or about 0.1 wt % to about 2.0 wt %. For example, the amount of phosphorus can be at least about 0.1 wt %, or at least about 0.25 wt %, or at least about 0.5 wt %, or at least about 0.75 wt %, or at least about 1.0 wt %. Additionally or alternately, the amount of phosphorus can be about 10 wt % or less, or about 5 wt % or less, or about 2.0 wt % or less, or about 1.5 wt % or less, or about 1.2 wt % or less, or about 1.1 wt % or less, or about 1.0 wt % or less.

A catalyst composition can employ a transition metal-enhanced zeolite in its original crystalline form or after formulation into catalyst particles, such as by extrusion. A process for producing zeolite extrudates in the absence of a binder is disclosed in, for example, U.S. Pat. No. 4,582,815, the entire contents of which are incorporated herein by reference. Preferably, the transition metal can be incorporated after formulation of the zeolite (such as by extrusion) to form self-bound catalyst particles. Optionally, a self-bound catalyst can be steamed after extrusion. The terms “unbound” and “self-bound” are intended to be synonymous and mean that the present catalyst composition is free of any of the inorganic oxide binders, such as alumina or silica, frequently combined with zeolite catalysts to enhance their physical properties.

The transition metal-enhanced zeolite catalyst composition employed herein can further be characterized based on activity for hexane cracking, or Alpha value. Alpha value is a measure of the acid activity of a zeolite catalyst as compared with a standard silica-alumina catalyst. The alpha test is described in U.S. Pat. No. 3,354,078; in the Journal of Catalysis, Vol. 4, p. 527 (1965); Vol. 6, p. 278 (1966); and Vol. 61, p. 395 (1980), each incorporated herein by reference as to that description. The experimental conditions of the test used herein include a constant temperature of about 538° C. and a variable flow rate as described in detail in the Journal of Catalysis, Vol. 61, p. 395. Higher alpha values correspond with a more active cracking catalyst. For an oxygenate conversion catalyst, Alpha values of at least 15 can be suitable, with alpha values greater than 100 being preferred. In particular, the Alpha value can be about 15 to about 1000, or about 50 to about 1000, or about 100 to about 1000.

As an alternative to forming self-bound catalysts, zeolite crystals can be combined with a binder to form bound catalysts. Suitable binders for zeolite-based catalysts can include various inorganic oxides, such as silica, alumina, zirconia, titania, silica-alumina, cerium oxide, magnesium oxide, yttrium oxide, or combinations thereof. For catalysts including a binder, the catalyst can comprise at least about 10 wt % zeolite, or at least about 30 wt %, or at least about 50 wt %, such as up to about 90 wt % or more. Generally, a binder can be present in an amount between about 1 wt % and about 90 wt %, for example between about 5 wt % and about 40 wt % of a catalyst composition. In some aspects, the catalyst can include at least about 5 wt % binder, such as at least about 10 wt %, or at least about 20 wt %. Additionally or alternately, the catalyst can include about 90 wt % or less of binder, such as about 50 wt % or less, or about 40 wt % or less, or about 35 wt % or less. Combining the zeolite and the binder can generally be achieved, for example, by mulling an aqueous mixture of the zeolite and binder and then extruding the mixture into catalyst pellets. A process for producing zeolite extrudates using a silica binder is disclosed in, for example, U.S. Pat. No. 4,582,815. Optionally, a bound catalyst can be steamed after extrusion.

In some aspects, a binder can be used that is substantially free of alumina, such as a binder that is essentially free of alumina. In this description, a binder that is substantially free of alumina is defined as a binder than contains about 10 wt % alumina or less, such as about 7 wt % or less, or about 5 wt % or less, or about 3 wt % or less. A binder that is essentially free of alumina is defined as a binder that contains about 1 wt % or less, such as about 0.5 wt % or less, or about 0.1 wt % or less. In still other aspects, a binder can be used that contains no intentionally added alumina and/or that contains no alumina within conventional detection limits for determining the composition of the binder and/or the reagents for forming the binder. Although alumina is commonly used as a binder for zeolite catalysts, due in part to ease of formulation of alumina-bound catalysts, in some aspects the presence of alumina in the binder can reduce or inhibit the activity of a transition metal-enhanced zeolite for converting methanol to aromatics. For example, for a catalyst where the transition metal is incorporated into the catalyst after formulation of the bound catalyst (such as by extrusion), the transition metal may have an affinity for exposed alumina surfaces relative to exposed zeolite surfaces, leading to increased initial deposition and/or migration of transition metal to regions of the bound catalyst with an alumina surface in favor of regions with a zeolite surface. Additionally or alternately, alumina-bound catalysts can tend to have low micropore surface area, meaning that the amount of available zeolite surface available for receiving a transition metal may be undesirably low.

As an example of forming a bound catalyst, the following procedure describes a representative method for forming silica bound ZSM-5 catalyst particles. ZSM-5 crystal and a silica binder, such as an Ultrasil silica binder, can be added to a mixer and mulled. Additional deionized water can be added during mulling to achieve a desired solids content for extrusion. Optionally, a caustic solution can also be added to the mixture and mulled. The mixture can then be extruded into a desired shape, such as 1/10″ quadralobes. The extrudates can be dried overnight at about 250° F. (121° C.) and then calcined in nitrogen for about 3 hours at about 1000° F. (538° C.). The extrudates can then be exchanged twice with an about IN solution of ammonium nitrate. The exchanged crystal can be dried overnight at about 250° F. (121° C.) and then calcined in air for about 3 hours at about 1000° F. (538° C.). This results in a silica bound catalyst. Based on the exchange with ammonium nitrate and subsequent calcinations in air, the ZSM-5 crystals in such a bound catalyst can correspond to ZSM-5 with primarily hydrogen atoms at the ion exchange sites in the zeolite. Thus, such a bound catalyst is sometimes described as being a bound catalyst that includes H-ZSM-5.

To form a transition metal-enhanced catalyst, a bound catalyst can be impregnated via incipient wetness with a solution containing the desired metal for impregnation, such as Zn or P. The impregnated crystal can then be dried overnight at about 250° F. (121° C.), followed by calcination in air for about 3 hours at about 1000° F. (538° C.). More generally, a transition metal can be incorporated into the zeolitic catalyst at any convenient time, such as before or after ion exchange to form H-form crystals, or before or after formation of a bound extrudate. In some aspects that are preferred from a standpoint of facilitating manufacture of a bound zeolite catalyst, the transition metal can be incorporated into the bound catalyst (such as by impregnation or ion exchange) after formation of the bound catalyst by extrusion or another convenient method.

Example of Reaction System Configuration

FIG. 1 shows an example of a reaction system configuration for performing oxygenate conversion to form a naphtha boiling range product. It is noted that the reactors shown in FIG. 1 are depicted as fixed bed, downflow reactors (such as trickle-bed reactors) for convenience. It is understood that any or all of the reactors shown in FIG. 1 can alternatively be moving bed reactors and/or fluidized bed reactors. In FIG. 1, a feed 105 can correspond to an oxygenate-containing feed. In a particular example, feed 105 can correspond to 96 wt % methanol and 4 wt % water. A second feed 106 can correspond to an olefin-containing feed. Optionally, oxygenate feed 105 can be introduced into a reactor as a plurality of input flows, such as a first input flow containing a mixture of methanol and water and a second input flow containing a mixture of nitrogen and hydrogen. Optionally, oxygenate feed 105 and olefinic feed 106 can be combined prior to entering the reactor 110.

The feed 105 (or alternatively a combination of oxygenate feed 105 and olefinic feed 106) can optionally be introduced into an initial dehydration reactor 110. Initial dehydration reactor 110 can include an acidic catalyst, such as an acidic alumina catalyst, that can facilitate an equilibrium reaction between methanol, water, and dimethyl ether. This can result in production of an effluent 115 that includes both methanol and dimethyl ether. Those of skill in the art will recognize that dimethyl ether and methanol can often be used in similar manners when performing an oxygenate conversion reaction. The dehydration of methanol to form dimethyl ether is highly exothermic. By performing an initial dehydration, the amount of heat generated in the conversion reactor(s) can be reduced, which can allow for improved temperature control in the conversion reactor. Optionally, a portion of the oxygenate feed 105 can bypass the dehydration reactor and can be input directly into conversion reactor 120. In aspects where other oxygenates are used as a feed, such as C₂₊alcohols or larger ethers, dehydration reactor can be omitted so that feed 105 (or a combination of oxygenate feed 105 and olefinic feed 106) is an input flow for conversion reactor 120.

The oxygenate feed 105 and olefinic feed 106 (and/or the effluent 115 containing both dimethyl ether and methanol) are then passed into conversion reactor 120. The input to conversion reactor 120 can be exposed to a conversion catalyst under effective conditions for forming a conversion effluent 125. The conversion effluent 125 can then be separated, such as by using a 3 phase separator 130. One phase generated by separator 130 can be an aqueous phase 133 that includes a substantial majority of the water present within the conversion effluent 125. Another phase generated by separator 130 can correspond to a hydrocarbon liquid product 137. The hydrocarbon liquid product can correspond to naphtha boiling range compounds formed during the conversion reaction. Optionally, the hydrocarbon liquid product can include a portion of hydrocarbon-like compounds that include one or more heteroatoms, such as oxygen, sulfur, nitrogen, and/or other heteroatoms that are commonly found in petroleum or bio-derived feeds.

A third phase generated by separator 130 can correspond to a hydrocarbon gas product 135. The hydrocarbon gas product 135 can include C⁴⁻compounds corresponding to light paraffins and light olefins. Optionally, a recycle portion 122 of hydrocarbon gas product 135 can be recycled as part of the input flows to conversion reactor 120. In some configurations where the amount of recycle portion 122 is sufficiently large, a bleed or waste flow (not shown) can also be present to reduce or minimize the build-up of C⁴⁻paraffins in conversion reactor 120.

EXAMPLE 1 Methanol Conversion Using Zeolitic Catalysts

Various conversion catalysts were tested in an isothermal fixed-bed reactor without recycle. In this example, the feed corresponded to 100 wt % methanol. The feed was exposed to conversion catalyst at a temperature of 450° C., a pressure of 15 psig, and a weight hourly space velocity of 2 hr⁻¹.

The ZSM-5 conversion catalysts used in this example were based on small crystal, self-bound MFI framework (ZSM-5) zeolite. The ZSM-5 had a silicon to aluminum ratio of 20 to 40 and an Alpha value of at least 100. For the catalyst with added Zn, after making an H-form extrudate of the self-bound zeolite, Zn was added via aqueous impregnation of Zn(NO₃)₂.

FIG. 2 shows the aromatics yield from conversion of a methanol feed versus the amount of feed processed for various types of zeolitic catalysts. The catalysts include self-bound ZSM-48; yttria bound ZSM-48 (65:35 zeolite to binder ratio); alumina bound ITQ-13 (65:35 zeolite to binder ratio); self-bound ZSM-5; and self-bound ZSM-5 with 0.5 wt % Zn supported on the catalyst. The ZSM-5 catalysts were similar to the catalyst used in Example 1. As shown in FIG. 2, for all of the zeolitic catalysts, the selectivity for aromatics formation starts out at an initial level for each catalyst and then steadily declines as greater amounts of feed are processed. FIG. 2 also shows that addition of 0.5 wt % Zn to a catalyst can increase the initial level of aromatics selectivity, but the decline in aromatics selectivity with exposure to feed remains similar.

EXAMPLE 2 Conversion of Methanol and Olefins with Zeolitic Catalyst Having Supported Transition Metal

A self-bound ZSM-5 catalyst similar to the catalyst in Example 1 was impregnated with zinc nitrate via incipient wetness to form a 1 wt % Zn/ZSM-5 catalyst. As shown in FIG. 2, addition of Zn to a zeolitic catalyst can improve the initial aromatic selectivity of a catalyst. The 1 wt % Zn/ZSM-5 catalyst was exposed to two types of feeds to determine the impact of addition of olefins in the presence of a metal-enhanced zeolitic catalyst. A first type of feed corresponded to 100 wt % methanol, similar to Example 1. A second type of feed corresponded to about 70 wt % methanol and about 30 wt % of 1-pentene. This corresponded to a feed with a oxygenate to olefin molar ratio of about 5. In this example, the feeds were exposed to the conversion catalyst in an isothermal reactor at a temperature of 450° C., a pressure of 90 psig, and a weight hourly space velocity of 2 hr⁻¹. The data shown in FIG. 3 corresponds to data taken after 46 hours of exposure of each feed to the catalyst, or at roughly 90 grams of feed per gram of catalyst.

FIG. 3 shows the product distribution for the hydrocarbon products generated from conversion of the methanol and the methanol/1-pentene feeds. Substantially all of the methanol and/or dimethyl ether in the feed was converted under these conditions. Due to the presence of 1 wt % Zn on the zeolitic catalyst, conversion of the methanol feed resulted in formation of at least ˜15 wt % (relative to the total hydrocarbon product) of carbon oxides (CO, CO₂) and methane. By contrast, the feed including both olefins and methanol had a carbon oxide plus methane yield of less than 4 wt %. This substantial reduction in the amount of methane and carbon oxides produced when using a combined feed of methanol and olefins resulted in a corresponding increase in aromatics production (i.e., aromatics selectivity). Although the ratio of aromatics versus paraffins or the ratio for aromatics versus olefins appeared to be unchanged when using the combined oxygenate plus olefins feed, the ability to avoid formation of methane and carbon oxides appeared to increase aromatics selectivity by more than 5 wt % (relative to the total hydrocarbon yield). Thus, it appears that using a combined feed of oxygenates and olefins provides an unexpected shift upward in aromatic selectivity for a conversion catalyst that includes a supported metal, as compared with the aromatic selectivity for a feed including only oxygenates.

EXAMPLE 3 Conversion of Methanol and Olefins with Zeolitic Catalyst (No Metal)

The decline in selectivity to carbon oxides and methane with exposure to feed was further investigated using the self-bound ZSM-5 catalyst without an additional supported metal. A feed corresponding to 100 wt % methanol was processed over self-bound ZSM-5 catalyst in an isothermal reactor at a total pressure of 15 psig (˜100 kPag) and a space velocity of 2 hr⁻¹. The temperature for the conversion was 325° C., 350° C., or 375° C. FIG. 4 shows results from conversion of the methanol feed over the conversion catalyst at the three temperatures. In FIG. 4, the initial portion of the conversion reaction was performed at 350° C., followed by a reduction to 325° C., and then the temperature was increased to 375° C.

FIG. 4 shows the relative yield of paraffins, aromatics, olefins, and “other” in comparison with the total converted hydrocarbon yield, as determined by gas chromatography. Similar to FIG. 2, the selectivity of the ZSM-5 catalyst for aromatics starts at an initial higher level and then declines with increasing feed exposure. This trend is visible at 350° C., and resumes when the conversion temperature is returned to 375° C. Additionally, while the trend of decreasing aromatics selectivity is visible at 325° C., the selectivity for aromatics formation is further shifted to lower values at this temperature. This is believed to be due to a substantial portion of (unconverted) methanol and/or dimethyl ether being present in the total hydrocarbon product. In other words, for the methanol/dimethyl ether that is converted at 325° C., it is believed that the selectivity remains the same, but the presence of a substantial amount of unconverted methanol/dimethyl ether in the total hydrocarbon product means that the yield of aromatics is lower.

FIG. 5 shows results from performing a similar conversion reaction using the ZSM-5 catalyst, but with a feed that included about 70 wt % methanol and about 30 wt % 1-pentene (similar to the oxygenate plus olefin feed used in Example 2). The combined methanol/1-pentene feed was converted under the same conditions as the methanol feed in FIG. 5, including using a sequence of conversion temperatures of 350° C., followed by 325° C., followed by 375° C. As shown in FIG. 5, the addition of 1-pentene reduced or minimized the difference in aromatics selectivity at 325° C. However, the addition of 1-pentene did not appear to impact the overall aromatics selectivity at 350° C. or 375° C. relative to the aromatics selectivity for the methanol feed. Thus, for a zeolitic catalyst without a supported metal, addition of olefins to the feed did not appear to modify overall aromatics selectivity at temperatures that were high enough to allow for full conversion of oxygenates in the feed.

Without being bound by any particular theory, it is believed that in addition to undergoing conversion, the olefins in the feed can also assist with conversion of oxygenates in the feed. This is demonstrated in FIG. 6, which shows the distribution of hydrocarbon products (including “products” corresponding to unreacted methanol/dimethyl ether) from the conversion runs in FIGS. 4 and 5. In FIG. 6, the left bar for each product type corresponds to the methanol feed, while the right bar corresponds to the methanol plus olefin feed. As shown in FIG. 6, exposing the methanol feed to the conversion catalyst at 325° C. resulted in a substantial amount of methanol/dimethyl ether that was not converted. By contrast, only a minimal amount of methanol was unconverted for the feed including both methanol and 1-pentene.

FIG. 7 shows additional details for product yields from the conversion reactions corresponding to FIGS. 4 and 5 at each of the studied temperatures (325° C., 350° C., and 375° C.).

FIG. 7 includes data for both the methanol feed and the mixed methanol plus 1-pentene feed. For each product type shown in FIG. 7 (water, C5+, C3-C4, C1-C2), the left three bars correspond to the methanol feed, while the right three bars correspond to the methanol plus olefin feed. Each series of three bars corresponds to 325° C., 350° C., and 375° C. from left to right. For the yields in FIG. 7, any unreacted methanol/dimethyl ether is not included as part of the yield calculation. As shown in FIG. 7, addition of olefins to the methanol feed resulted in a lower amount of C₁ and/or C₂ production (such as production of methane). The amount of water generated was reduced, but that is at least partially explained by the reduction in oxygenates in the feed. Similarly, the apparent increase in olefins for the feed including both methanol and olefins may be related to the additional olefins that were present in the feed.

EXAMPLE 4 Reduction in Durene Formation

FIG. 8 shows additional details regarding the types of aromatics that were produced during conversion of the methanol feed and the methanol/1-pentene feed using the ZSM-5 catalyst (no metal). The results in FIG. 8 were obtained at the reaction conditions listed in conjunction with FIGS. 4 and 5. In FIG. 8, “A6” corresponds to an aromatic compound that includes 6 carbon atoms (i.e., benzene). In other words, “A6” refers to a C₆ aromatic compound. The “A7”, “A8”, and “A9” columns are believed to correspond to various alkyl-substituted benzenes. The “A10” and “A11” columns can potentially include both alkyl substituted benzenes, naphthalene, and substituted naphthalenes. For each product type shown in FIG. 8, the left three bars correspond to the methanol feed, while the right three bars correspond to the methanol plus olefin feed. Each series of three bars corresponds to 325° C., 350° C., and 375° C. from left to right. As shown in FIG. 8, addition of olefins to the methanol feed resulted in a substantial reduction in the amount of durene produced at all temperatures (325° C., 350° C., 375° C.). Durene corresponds to 1, 2, 4, 5 tetramethylbenzene, and therefore represents a substantial portion of the aromatics present in the “A10” aromatics column. Durene is a compound that can crystallize at relatively low temperatures, and that can potentially affect gasoline performance and appearance. The ability to suppress durene formation while maintaining a similar overall selectivity for aromatics formation is another unexpected advantage of using a combined feed of oxygenates and olefins as the conversion feed.

FIG. 9 shows that a similar reduction in durene formation was achieved when performing conversion of the methanol/1-pentene feed using a 1 wt % Zn/ZSM-5 catalyst. The results in FIG. 9 were generated by conversion of a methanol feed or a methanol/1-pentene feed in an isothermal reactor at a temperature of 450° C., a pressure of 90 psig, and a WHSV of 2 hr⁻¹. In FIG. 9, the left bar for each product type corresponds to the methanol plus olefin feed, while the right bar corresponds to the methanol feed. It is noted that the increased temperature used for the results in FIG. 9 also contributed to reducing the amount of durene production. However, the benefit of using a combined oxygenate/olefin feed for reducing durene production is still evident in FIG. 9. Based on FIG. 9, addition of olefins when performing oxygenate conversion can allow for formation of a conversion effluent that includes less than 10 wt % C₁₀ aromatics relative to the total weight of hydrocarbons in the conversion effluent. Additionally or alternately, less than 10 wt % of the C₁₀ aromatics (or less than 5 wt %) can correspond to durene.

EXAMPLE 5 Conversion Using Non-MFI Zeolitic Catalysts

The results in Example 1 show the relatively short processing lifetimes for non-MFI zeolitic catalysts for conversion of oxygenates to olefins and/or aromatics. In particular, as the amount of oxygenate exposed to the catalyst increases, such as to more than about 100 g oxygenate per gram of catalyst, the selectivity of non-MFI catalysts for oxygenate conversion rapidly falls off. It has been unexpectedly discovered that using a combined oxygenate and olefin feed can reduce or minimize this fall off in conversion activity for non-MFI zeolitic catalysts.

FIG. 10 shows results from conversion of a methanol feed (similar to the feed in Example 1) for a ZSM-48 catalyst. The ZSM-48 catalyst did not include an additional metal supported on the catalyst. The conversion reaction was performed in an isothermal reactor at 350° C., 15 psig (˜100 kPag), and a WHSV of about 2 hr⁻¹. A similar conversion reaction using the ZSM-48 catalyst was also performed using the combined methanol/1-pentene feed described above. The results from conversion of the combined methanol/1-pentene feed are shown in FIG. 11.

FIG. 10 appears to show that a substantial portion of the loss in selectivity for olefins formation when using a methanol feed is due to a loss in ability to convert the methanol feed. For example, the amount of unconverted methanol corresponds to about 20 wt % of the feed at an exposure of only 100 grams of methanol per gram of catalyst. By contrast, when an olefin is also present in the feed, FIG. 11 shows that at least 90 wt % conversion of the feed is maintained until at least an exposure of 150 grams of feed per gram of catalyst. Based on extrapolation, it further appears that roughly 90 wt % conversion of feed could be maintained until at least an exposure of 200 grams of feed per gram of catalyst. Surprisingly, the aromatics selectivity in FIGS. 10 and 11 appears to be similar, even as FIG. 10 shows substantially lower amounts of overall feed conversion. Thus, for non-MFI zeolitic catalysts, inclusion of olefins in an oxygenate feed appears to provide a benefit by allowing for increased feed conversion to olefins, while maintaining a somewhat similar selectivity for aromatics formation. Based on the data in FIG. 11, addition of olefins can allow for operation of a conversion reactor at an average catalyst exposure to feed of between about 50 grams of feed per gram of catalyst to about 200 grams of feed per gram of catalyst, (or about 50 grams to about 150 grams, or about 75 grams to about 200 grams, or about 75 grams to about 175 grams, or about 75 grams to about 150 grams, or about 100 grams to about 200 grams), while still maintaining at least 85 wt % conversion of oxygenate in the feed, or at least 90 wt %. In other words, the oxygenate in the conversion effluent can be 15 wt % or less of the total hydrocarbon product, or 10 wt % or less.

Additional Embodiments

Embodiment 1. A method for forming a naphtha composition, comprising: exposing a feed comprising oxygenates and olefins to a conversion catalyst at an average reaction temperature of about 300° C. to about 550° C., a total pressure of about 10 psig (˜70 kPag) to about 400 psig (˜2700 kPag), and a WHSV of 0.1 hr⁻¹ to 20.0 hr⁻¹ to form a converted effluent comprising a naphtha boiling range fraction having an octane rating of at least 80, the converted effluent further comprising less than 6.0 wt % combined of CO, CO₂, and CH₄ relative to a total weight of hydrocarbons in the converted effluent, the feed having a molar ratio of oxygenates to olefins of about 1 to about 20, wherein the conversion catalyst comprises at least 10 wt % of a zeolite having MFI framework structure, the zeolite having a silicon to aluminum ratio of 10 to 200 (or 20 to 40) and an Alpha value of at least 5 (or at least 15), the conversion catalyst further comprising 0.1 wt % to 3.0 wt % of a transition metal supported on the conversion catalyst.

Embodiment 2. The method of Embodiment 1, wherein the naphtha boiling range fraction comprises an octane rating of at least 90 (or at least 93, or at least 97) and at least about 40 wt % aromatics relative to a weight of the naphtha boiling range fraction.

Embodiment 3. The method of any of the above embodiments, wherein the average reaction temperature is at least about 400° C., or at least about 450° C.

Embodiment 4. The method of any of the above embodiments, wherein the oxygenates comprises methanol, the conversion catalyst comprising an average catalyst exposure time of 1 gram to 2000 grams of oxygenate per gram of catalyst, or 1 gram to 200 grams, or 400 grams to 1500 grams.

Embodiment 5. A method for forming a naphtha composition, comprising: exposing a feed comprising oxygenates and olefins to a conversion catalyst at an average reaction temperature of about 300° C. to about 550° C., a total pressure of about 10 psig (˜70 kPag) to about 400 psig (˜2700 kPag), and a WHSV of 0.1 hr⁻¹ to 20.0 hr⁻¹ to form a converted effluent comprising a naphtha boiling range fraction and further comprising at least about 30 wt % olefins and less than 15 wt % oxygenate relative to a total weight of hydrocarbons in the converted effluent, the feed having a molar ratio of oxygenates to olefins of 1 to 20, wherein the conversion catalyst comprises at least 10 wt % of a 10-member ring or 12-member zeolite having a framework structure different from WI (optionally different from WI or MEL) framework structure, the zeolite having a silicon to aluminum ratio of 10 to 200 (or 20 to 40) and an Alpha value of at least 5 (or at least 15), the conversion catalyst further comprising an average catalyst exposure time of 25 grams to 200 grams of oxygenate per gram of catalyst, the conversion catalyst optionally further comprising 0.1 wt % to 3.0 wt % of a transition metal (preferably Zn) supported on the conversion catalyst.

Embodiment 6. The method of Embodiment 5, wherein the oxygenate comprises methanol, or wherein the conversion catalyst comprises an average catalyst exposure time of 50 grams to 200 grams of methanol per gram of catalyst, (or 25 grams to 180 grams, or 50 grams to 180 grams, or 50 grams to 150 grams, or 100 grams to 200 grams); or a combination thereof.

Embodiment 7. The method of Embodiment 5 or 6, wherein the conversion catalyst comprises at least 10 wt % of a zeolite having a framework structure of MRE (ZSM-48), MTW, TON, MTT, MFS, or a combination thereof.

Embodiment 8. The method of any of Embodiments 5 to 7, wherein exposing the feed comprising oxygenates to a conversion catalyst comprises exposing the feed comprising oxygenate to the conversion catalyst in a fluidized bed, a moving bed, a riser reactor, or a combination thereof, the conversion catalyst being withdrawn and regenerated at a rate corresponding to regeneration of 0.3 wt % to 3.0 wt % of catalyst per 1 g of oxygenate exposed to a g of conversion catalyst (optionally 1.5 wt % to 3.0 wt %).

Embodiment 9. The method of any of the above embodiments, wherein the 0.1 wt % to 3.0 wt % of transition metal comprises 0.1 wt % to 3.0 wt % of Zn, or 0.5 wt % to 1.5 wt % Zn.

Embodiment 10. The method of any of the above embodiments, wherein the conversion catalyst further comprises phosphorus supported on the conversion catalyst, a molar ratio of phosphorus to zinc on the conversion catalyst optionally being 1.5 to 3.0.

Embodiment 11. The method of any of the above embodiments, wherein the feed comprises a molar ratio of oxygenates to olefins of 10 or less, or 6.0 or less.

Embodiment 12. The method of any of the above embodiments, wherein the feed comprising oxygenates and olefins comprises a first feedstock comprising at least a portion of the oxygenates and a second feedstock comprising at least a portion of the olefins, the first feedstock and the second feedstock being combined after entering a reactor containing the conversion catalyst.

Embodiment 13. The method of any of the above embodiments, a) wherein the feed comprises about 30 wt % to about 95 wt % of oxygenates, about 5 wt % to about 40 wt % of olefins, or a combination thereof; b) wherein the feed comprises at least about 20 wt % to about 60 wt % of components different from oxygenates and olefins, or about 40 wt % to 60 wt %; or c) a combination of a) and b).

Embodiment 14. An oxygenate conversion effluent comprising, relative to a total weight of hydrocarbons in the conversion effluent, at least 40 wt % aromatics, less than 6.0 wt % combined of CO, CO₂, and CH₄, and less than 10 wt % olefins, a naphtha boiling range portion of the conversion effluent having an octane rating of at least 90, wherein less than 10 wt % of the aromatics comprise C₁₀ aromatics relative to a total weight of the aromatics, and wherein less than 10 wt % of the C₁₀ aromatics comprise durene relative to a total weight of the C₁₀ aromatics.

Embodiment 15. The oxygenate conversion effluent of Embodiment 14, wherein the oxygenate conversion effluent comprises less than 5.0 wt % combined of CO, CO₂, and CH₄, or wherein less than 5 wt % of the C₁₀ aromatics comprise durene relative to a total weight of the C₁₀ aromatics, or a combination thereof.

Embodiment 16. A conversion effluent made according to any of Embodiments 1-13.

While the present invention has been described and illustrated by reference to particular embodiments, those of ordinary skill in the art will appreciate that the invention lends itself to variations not necessarily illustrated herein. For this reason, then, reference should be made solely to the appended claims for purposes of determining the true scope of the present invention. 

1. A method for forming a naphtha composition, comprising: exposing a feed comprising oxygenates and olefins to a conversion catalyst at an average reaction temperature of about 300° C. to about 550° C., a total pressure of about 10 psig (˜70 kPag) to about 400 psig (˜2700 kPag), and a WHSV of 0.1 hr⁻¹ to 20.0 hr⁻¹ to form a converted effluent comprising a naphtha boiling range fraction having an octane rating of at least 80, the converted effluent further comprising less than 6.0 wt % combined of CO, CO₂, and CH₄ relative to a total weight of hydrocarbons in the converted effluent, the feed having a molar ratio of oxygenates to olefins of about 1 to about 20, wherein the conversion catalyst comprises at least 10 wt % of a zeolite having MFI framework structure, the zeolite having a silicon to aluminum ratio of 10 to 200 and an Alpha value of at least 5, the conversion catalyst further comprising 0.1 wt % to 3.0 wt % of a transition metal supported on the conversion catalyst.
 2. The method of claim 1, wherein the naphtha boiling range fraction comprises an octane rating of at least 90 and at least about 40 wt % aromatics relative to a weight of the naphtha boiling range fraction.
 3. The method of claim 1, wherein the average reaction temperature is at least about 400° C.
 4. The method of claim 1, wherein the 0.1 wt % to 3.0 wt % of transition metal comprises 0.1 wt % to 3.0 wt % of Zn.
 5. The method of claim 1, wherein the conversion catalyst further comprises phosphorus supported on the conversion catalyst.
 6. The method of claim 1, wherein the conversion catalyst comprises 0.5 wt % to 1.5 wt % Zn supported on the conversion catalyst.
 7. The method of claim 6, wherein the conversion catalyst further comprises phosphorus, a molar ratio of phosphorus to zinc on the conversion catalyst being 1.5 to 3.0.
 8. The method of claim 1, wherein the feed comprises a molar ratio of oxygenates to olefins of 10 or less.
 9. The method of claim 1, wherein the feed comprising oxygenates and olefins comprises a first feedstock comprising at least a portion of the oxygenates and a second feedstock comprising at least a portion of the olefins, the first feedstock and the second feedstock being combined after entering a reactor containing the conversion catalyst.
 10. The method of claim 1, wherein the feed comprises about 30 wt % to about 95 wt % of oxygenates, about 5 wt % to about 40 wt % of olefins, or a combination thereof.
 11. The method of claim 1, wherein the feed comprises at least about 20 wt % to about 60 wt % of components different from oxygenates and olefins.
 12. The method of claim 1, wherein the oxygenates comprises methanol, the conversion catalyst comprising an average catalyst exposure time of 1 grams to 2000 grams of oxygenate per gram of catalyst.
 13. A method for forming a naphtha composition, comprising: exposing a feed comprising oxygenates and olefins to a conversion catalyst at an average reaction temperature of about 300° C. to about 550° C., a total pressure of about 10 psig (˜70 kPag) to about 400 psig (˜2700 kPag), and a WHSV of 0.1 hr⁻¹ to 20.0 hr⁻¹ to form a converted effluent comprising a naphtha boiling range fraction and further comprising at least about 30 wt % olefins and less than 15 wt % oxygenate relative to a total weight of hydrocarbons in the converted effluent, the feed having a molar ratio of oxygenates to olefins of 1 to 20, wherein the conversion catalyst comprises at least 10 wt % of a 10-member ring or 12-member zeolite having a framework structure different from MFI framework structure, the zeolite having a silicon to aluminum ratio of 10 to 200 and an Alpha value of at least 5, the conversion catalyst further comprising an average catalyst exposure time of 25 grams to 200 grams of oxygenate per gram of catalyst.
 14. The method of claim 13, wherein the oxygenate comprises methanol, and wherein the conversion catalyst comprises an average catalyst exposure time of 50 grams to 180 grams of methanol per gram of catalyst.
 15. The method of claim 13, wherein the conversion catalyst comprises at least 10 wt % of a zeolite having a framework structure of MRE (ZSM-48), MTW, TON, MTT, MFS, or a combination thereof.
 16. The method of claim 13, wherein the conversion catalyst further comprises 0.1 wt % to 3.0 wt % of Zn supported on the conversion catalyst.
 17. The method of claim 13, wherein the conversion catalyst further comprises phosphorus supported on the conversion catalyst.
 18. The method of claim 13, wherein exposing the feed comprising oxygenates to a conversion catalyst comprises exposing the feed comprising oxygenate to the conversion catalyst in a fluidized bed, a moving bed, a riser reactor, or a combination thereof, the conversion catalyst being withdrawn and regenerated at a rate corresponding to regeneration of 0.3 wt % to 3.0 wt % of catalyst per 1 g of oxygenate exposed to a g of conversion catalyst.
 19. An oxygenate conversion effluent comprising, relative to a total weight of hydrocarbons in the conversion effluent, at least 40 wt % aromatics, less than 6.0 wt % combined of CO, CO₂, and CH₄, and less than 10 wt % olefins, a naphtha boiling range portion of the conversion effluent having an octane rating of at least 90, wherein less than 10 wt % of the aromatics comprise C₁₀ aromatics relative to a total weight of the aromatics, and wherein less than 10 wt % of the C₁₀ aromatics comprise durene relative to a total weight of the C₁₀ aromatics.
 20. The oxygenate conversion effluent of claim 16, wherein the oxygenate conversion effluent comprises less than 5.0 wt % combined of CO, CO₂, and CH₄, or wherein less than 5 wt % of the C₁₀ aromatics comprise durene relative to a total weight of the C₁₀ aromatics, or a combination thereof. 